Process for improving the octane number of naphthas



H. HENNIG Aug. 2, 1960 PROCESS FOR IMPROVING THE OCTANE NUMBER OF NAPIV-ITHAS Filed Dec. l2, 1957 United States PatetOfice Patented Aug. 2., 1960 PROCESS FOR llVIPROVING THE OCTANE NUlVIBER F NAPHTHAS Harvey Hennig, Crystal Lake, Ill., assignor to The Pure Oil Company, Chicago, Ill., a corporation of Ohio Filed Dec. 12, 1957, Ser. No. 702,442

3 Claims. (Cl. 208-65) This invention relates to the processing of full-boilingrange, straight-run gasolines to produce high-octanenumber blending stocks. It is more specifically concerned with the integration of a plurality of unit petroleum refining processes to provide a combination operation for the upgrading of straight-run gasoline.

According to this invention, the unit processes of catalytic refining, catalytic demethylation and catalytic hydroisornerization are combined with adjunct fractionation processes, including solvent refining, adsorption and ythermal fractional distillation, to provide an integrated scheme for the efficient upgrading of straight-run gasolines to produce high yields of a high-octane-nurnber blending stock.

Motor fuels required for the efficient operation of modern high-output, high-speed, spark-ignited, internal combustion engines are high-octane-number formulations produced by blending a number of selected blending stocks. Quality as well as progressive competitors have necessitated that refiners of petroleum products evaluate 'the inter-relation of existing rening processes and make appropriate changes to meet the exigencies of the situation. In processing normally liquid hydrocarbons boil- .ing inl the gasoline range, reforming has played an important part. When straight-run gasolines are used as `feedstocks, constituents such as the low-boiling saturated hydrocarbons are not efficiently processed for maximum improvement in octane number. `It is more desirable in processing these straight-run gasolines to high octane levels to use several processes designed to be more selective to` a `specific reaction among the number of reactions `occurring `in conventional reforming.

It is, therefore, the primary object of this invention toprovide a combination of processes centering around reforming `for octane-number improvement ofstraightrun gasolines. It is another object of this invention to 'provide a combination of conventional refining processes 'integrated to produce high-octane-number blending stocks from straight-run gasoline. These and other objects will become more apparent from the following detailed description of this invention.

VThe single figure shows a simplified flow sheet of the process of this invention. t. It is known `that full-boiling-range, straight-run gasolines ,contain A,normaland branch-chain paraffins, naphthenes, and aromatics. `Of these major types of constituents, the .branch-chainparafiinic hydrocarbons and aromatics, in general, have relatively high octane numbers and are ,suitable for direct use in motor fuel. Naphthenes, even though they have intermediate yoctane numbers, can be converted to relatively high octane number aromatics by components. the bottom` of extraction tower 21 through line 23, to

n-heptane will be marginal in some instances.

reforming, thereby increasing the octane number of a refinery gasoline pool. The low-boiling, straight-chain parains (pentane, hexane and heptane) can be converted to branch-chain parafins which have relatively high octane numbers. However, octane, nonane, and other heavy straight-chain parafiins are not satisfactory for these isomerization processes. Even the isomerization of While it is possible to hydrocrack these heavier straight-chain parafiinic hydrocarbons in reforming operations, conditions must be so severe that large losses as gas and coke result. Accordingly, this approach is undesirable, even though commercially practiced. According to this invention, a process for the efficient conversion of full-boiling, straight-run gasolines to high-octane blending stocks at exceptionally high yield has been developed.

Referring to the simplified flow diagram in the single figure, obvious auxiliary equipment such as valves, pumps, heat exchangers, etc., have been omitted. for simplicity. Also, major process steps are shown schematically. In the process of this invention, debutanized, straight-run gasoline having a true end-point within the range of 370 to 420 F., having been desulfurized in a conventional manner in feed preparation unit 10, enters through line 11 and is Vdistilled in fractionator 12 into light and heavy fractions. The heavy fraction has a true boiling range of 193 F. to S70-420 F., contains naphthenes,

"some aromatics and heavy parat-linie hydrocarbons, and flows to catalytic reformer 13 through line 14. Reformer 13 is operated at conditions such that the naphthenes are substantially dehydrogenated, and thereby converted to aromatic hydrocarbons, and part of the heavy parafiins 'are hydrocracked. Conditions, however, are maintained 'at a mild conversion level so that the gas and coke make are minimized and high yields are obtained.

Efiiuent from catalytic reformer 13, after being freed of hydrogen gas and debutanized by suitable means not shown, is conducted through line 15 to raffinate fractionator 16. This tower divides the stream into heavy and light fractions, the heavy, predominantly aromatic fractions boiling above about 290 F. being withdrawn `as product through line 17. The light reformate, consystem, the aromatics are extracted from the light reformate and `after separation from the solvent, they are withdrawnV as product through line 20. Rafinate from extraction system 19 is conducted to raffinate fractionator 21 through line 22. In rafiinate fractionatorZl a separation is made between the high-octane isoheptanes (e.g., dimethyl pentanes) andV lighter components, and lowoctane isoheptanes (e.g., methyl hexanes) and heavier The heavier fraction is withdrawn from demethylation reactor 24. In this reactor terminal methyl groupsare selectively removed from the heavy parafiin molecules and some ring-breaking and demethylation of the smaller amounts of heavier naphthenes present takes place. Upon being returned to raflinate fractionator 21 through line 25, those components that have been demethylated sufficiently to form high-octane (lower boil ing) dimethyl pentanes, isohexanes, or lighter materials,

are separated and taken overhead with the lighter components of the raffinate through line 26. 'Those materials that are still heavier than desired are recycled through the demethylation reactor until they have become volatile enough to boil below about 193 F. If the naphthenes passing through the demethylation unit are of such a character that they do not readily thermally decompose to form paraflins, a small slip-stream, such as that indicated by broken line 27, is provided for recycling these naphthenes to reformer 13.

- The overhead stream from raflnate fractionator 21 containing only parainic and naphthenic compounds boiling belowY 193 F. joins with the overhead from naphtha fractionator 12, flowing through line 28, to form stream Z9 which passes to system 30 for separating branchedchain paraiinic hydrocarbons. In this system, normal paraffns are selectively removed from the stream, leaving only isomeric parailins 4and light naphthenes to pass overhead through line 31. These materials are removed as product. The straight-chain parainic hydrocarbons are passed through line 32 to catalytic isomerizationv reactor 33 where they are partially converted to branch-chain parafns. The efliuent from this reactor is returned to separation system 30 via line 34, the branch-chain parafnic hydrocarbons are removed as product, andV the straight-chain parafftnic hydrocarbons are recycled to isomerization reactor 33. Accordingly, the straight-chain l paraflinic hydrocarbons are recycled to extinction. Product streams 17, and 31 may be combined to form a high-octane-number lblending stock, or parts of the streams may be employed for selective blending of motor fuels.

In order to elfectuate the objectives of this invention, each of the component unit processes must perform a specific function. In order to function compatibly in the interrelated process, each unit process must be operated within a specific range of operating conditions as set forth in the following detailed discussion of each phase of the combination process of this invention. The function of the catalytic reforming section is to achieve maximum aromatization of the heavy naphtha fraction with a minimum of hydrocracking of octane and lighter paran hydrocarbons. Because of the superior yieldconversion relationship that is obtained employing platinum or palladium catalysts, this reforming phase is carried out in the presence of a reforming catalyst comprisp ing platinum or palladium supported on alumina, or on a siliceous cracking catalyst. Catalysts of this type are well known in the prior art and contain about 0.1 to 1.0% by weight of platinum or palladium on the selected support. Complete descriptions of such catalysts are contained in a number of U.S. patents, e.g., 2,478,916; 2,479,109; 2,550,531; 2,589,189; 2,705,329; and others. If desired, the catalysts can have a combined halogen (Od-0.8% of uorine or chlorine based on the support) incorporated therein.

These catalysts can be prepared in accordance with conventional techniques by impregnating the selected support, e.g., alumina, silica-alumina, silica-zirconia, silicaalumina-zirconia, etc., with an aqueous solutionV of chloroplatinic acid or chloropalladic acid, or with an aqueous solution of the ammonium salts of these acids. The impregnated support is dried at about 210-250 F. and reduced in a hydrogen atmosphere at a temperature within the range of about 200-250 F. (cf. Industrial and Engineering Chemistry, 45, 147, 1953). Operating conditions for carrying out the reforming steps are selected for producing optimum aromatization. Accordingly, the following' conditions are employed:

Temperature 900-930 F. Pressure Z50-350 p.s.i.g. Space velocity 1-2 liquid rate hourly space velocity. Hydrogen rate 4500-7000- fha/barrel of feed stock.

The solvent extraction process for lthe recovery of aromatic hydrocarbons, including benzene, toluene and xylene, from the C5-290 F. fraction obtained from the reformate is employed because conventional distillation cannot be used since aromatic and certain non-aromatic hydrocarbons tend to form azeotropes. The solvent used must have a high solubility for aromatics, com`bined with good selectivity. It must be stable lat the operating temperature employed and preferably has a high boiling point to permit the distillation of the aromatic hydrocarbons therefrom. The latter characteristic greatly reduces utility requirements. Specic solvents include glycols, polyglycols and acyl derivatives, etc., such as diethylene glycol and triethylene glycol, phenol, nitrobenzene, furfural, Chlorex and others. In addition, mixed solvents, such as those having the foregoing characteristics discused by Kalichevsky, ACS Monograph #76, Reinhold, can also be employed. Preferably, diethylene and triethylene glycols are utilized. In accordance with conventional practice, it is preferred to employ suiiicient amounts of water to permit adjusting the selectivity, solubility, and solvency of the solvent to the optimum for the individual case and at the same time provide a partial pressure of water vapor such that the aromatics can be recovered by stripping or distillation at moderate temperatures and pressures. Generally, 2.5 to 10 volumes of solvent per volume of water is used. In carrying out the extraction, a solvent/feed volume ratio of 10 to 30 to 1 is used employing extraction temperatures Vin the range of 110-300" F. The time of contacting should -be sufficient for attainment of essentially equilibrium conditions, and if necessary, plural-stage extraction should be used. Selection `of optimum conditions will depend upon the characteristics of the C5290 F. fraction being treated and can be readily determined experimentally (Vide U.S. Patent 2,803,685). The solvent extraction process is operated at a pressure sucient to maintain the system -in the liquid phase. Conventional extraction apparatus such as` sieve-tray, packed, or other type of towers can be used.

To promote demethylation of the feed to demethylation reactor 24, the reactor is preferably packed with a catalyst capable of relatively demethylating hydrocarbons, such as silica-alumina cracking catalyst impregnated with nickel molybdate, or with palladium and the catalyst is then subjected to a conditioning treatment. A silicaalumina cracking catalyst containing 75% by weight of silica and 25% by weight of alumina and impregnated with 10% Iby weight of nickel molybdate, then dried and pelleted into pellets having a diameter of about 1s-% inch, can be activated to a selective demethylation state as follows. The catalyst is oxidized with a mixture of oxygen and inert gas containing oxygen at a partial pressure of 30-500 mm. Hg, at a temperature of (a) The oxidized catalyst is contacted with a flowing v stream of substantially dry, hydrogen-containing gas at a temperature of about 60G-975 F. for a period suflicient to remove water, generated by the reduction reaction, from the catalyst bed;

(b) lThe catalyst is contacted further with dry hydrogen,

either by COlliiulling oW or shutting in the catalyst for -30 hours While temperature is `maintained at 900-973 F., and pressure at 0-600 p.s.i.g.; and (c) The reduced catalyst then is contacted with owling hydrogen containing about -30 mm. Hg water partial pressure, at a temperature of about 875 F. for a period of 15-30 hours.

4and/ or hydrogen therefrom.

(2) Oxidaton.-The catalyst is contacted with an oxygen-containing inert gas, containing oxygen at a partial pressure of 10-500 mm. Hg and Water vapor at a partial pressure of 0-90 mm. Hg, at a temperature of 825-975o F. for l-20 hours. Temperature, Water-vapor partial pressure, and length of treatment are critically interrelated in this step. At an oxidation temperature of 825 F., the Water-vapor partial pressure must be in the range of 20-90 mm. Hg and the length of treatment must be about 20 hours. At an oxidation temperature of 975 F., the Water-Vapor partial pressure must be in the range of 0-15 mm. Hg, and the length of treatment must be limited to about 1 hour.

(3) Purge.-Oxygen, combustion products, and moisture then are removed from the catalyst bed by purging with inert gas, or by evacuation to a subatmospheric pressure in the range of l0 mm. Hg, absolute.

s(4) Reduction.-The reducible constituents of the oxidized catalyst composition then are reduced to the greatest extent possible by passing a stream of dry, hydrogen-containing gas through the catalyst bed for L30 `hours at a temperature of S50-950 F. and a pressure of 15-750 p.s.i.g. While dry gas is preferred, water-vapor` partial pressures as high as l0 mm. Hg can be employed vwithout serious deleterious effect.

` Because the function of the catalytic .demethylation step is to remove a terminal methyl group from the heptane vand heavier parainic hydrocarbons to form parafns 4lighter than heptane, as well as to decyclize naphthenic rings and demethylate the resulting parafnic hydrocarbons, operating temperatures are selected which favor idemethylation over indiscriminate hydrocracking. Ac-

cordingly, the operating conditions employed in the de- `methylation step are as follows: i

Temperature F 620-680 Pressure p.s. i 300-750 Space velocity l-3 liz/hydrocarbon mol ratio 1-3/1 The separating system employed in processing the C5-193? F. overhead fractions obtained from the naphtha fractionator, andthe fractionator employed in separating 6 6 to 25 F. The alternative adsorption processes also can effect the substantially quantitative separation of straight-chain paraffinic hydrocarbons from branch-chain parainic hydrocarbons and naphthenes. If this alternative is employed as the fractionation technique, it is preferred that so-called molecular-sieve adsorbents be employed in this service. These adsorbents are available from the Linde Air Products Company and are composed of crystalline sodium and calcium alumino-silicates that have been heated to remove their waterof hydration. The resulting crystals are highly porous and the pores are of molecular dimensions, being only S15-20 billionths of an inch in diameter and of uniform size. A calcium alumino-silicate-type molecular sieve having a pore diameter of about 5 Angstroms will quantitatively separate the straight-chain parainic hydrocarbons from the branch-chain parains and naphthenes. -In U.S. Patent 2,306,610 is described the fractionation of mixtures of hydrocarbons employing molecular sieves of the crystalline zeolite type. Other adsorbents employed for selectively adsorbing straight-chain parafns from branchchain paraffin hydrocarbons or cyclic hydrocarbons inlclude activated coconut charcoal described in U.S. Patent 2,425,535, Saran charcoal prepared by the pyrolysis of polyvinylidene, and other similar adsorbents having small and uniform pore diameters sufcient to selectively Separate straight-chain paraiinic hydrocarbons from admixtiure with cyclic and branch-chain compounds. To effectuate this separation, pore diameters of about 5 Angstroms are necessary. i

The straight-chain paratlinic hydrocarbons which are recovered from the fractionation system 30 are subsequently converted to branch-chain paratnic hydrocarbons in a catalytic isomerization step carried out under conditions which will effect the isomerization Without concomitant `dernethylation and hydrocracking. Catalysts which will convert substantial amounts of the the raffinate stream obtained in the solvent extraction process into ,its lower-boiling and higher-boiling constituents, has the function of` separating normal `parains from branch-chain paraflns and light naphthenes for subsequent isomerization. This separation can be achieved by precision fractionation or selective adsorption. lf fractionation is used, conventional towers are employed to separate normal pentane and normal hexane from the branch-chain and naphthenic materials. If preferred, a superfractionation system can be used to separate the normal pentane and normal hexane constituents from the feed introduced into this fractionation system. In such an installation a plurality of bubble-tray towers ernploying bubble-trays is used. The number of trays, their diameter and spacing, as Well as reflux ratios and other details, are selected in order to separate stream con- Stituents with boiling point differentials of only about straight-chain parainic hydrocarbons to branch-chain hydrocarbons with excellent selectivity are the noble metal catalysts, such as platinum and palladium incorporated on refractory, mixed oxides, acidic-type carriers. About 0.1% to 1.0% by weight of the noble metal is included or incorporated in the mixed oxides support. These catalysts can be prepared as described above in connection with the noble metal catalysts as utilized in the reforming phase of the invention, preferably using 75-25 or -13 silica-alumina as the cracking catalyst support. High space velocities are preferred because they favor isomerization over hydrocracking, thereby achieving superior conversion-yield relationships. The following conditions are employed in the isomerization phase of this invention. p p

Temperature 650-800 Pressure -1000 p.s.i.g.

Space velocity 1-5 liquid rate hourly space velocity.

H2/hydrocarbon mol ratio 1-3/ 1.

To illustrate the advantages which are obtained in employing the combination process of this invention for the production of high-octane-number blending stocks, the combination of a catalytic reforming process employing a platinum catalyst with a solvent extraction unit employing diethylene glycol, and an isomerization process to make a high-octane-number blending material are compared with the improved integrated process of this invention. The results of this comparative study are tabulated in Table I in which case I represents the data obtained employing catalytic reforming, solvent extraction and isomerization, and case II is exemplary of the data obtained employing this invention.

Case I case 1r Straight-Run gasoline, percent v, (10D-380 F., ASTM boiling range from Illinois Crude). 193 F. EJ?. to isomerlzatiom )B3-400 F. to reforming Reforming: 1

Debutanlzed yield, percent v. of charge.. F-l octane, +3 cc. TEL Heavy reformate (29o-440 F. true honing point),

percent v. of charge Percent v. of Total S.R 2

Extraction: i

Extract (benzene, toluene, Xyl.)-

Percent Vv. of Reformate.- Percent v. of Total S.YR.

Rainate (C5-Cp Parans) to isome Total S.R 20

Recycle to reforming, v. percent Total S.R 33

Demethylation: 3

Fresh heavy raffinate reactor feed, v. percent Total 5! R 18 Recycle, to reactor, percent v. Total S.R 1g

Total Clt-193 F. liquid to isomerization. Isomerization: 4 y

Fresh feed, percent v. TotalS.R.

Product, percent v. Total S.Rr

ProdEuIclts, VOL percent of total B.R. and blending octanes, F-l-l-3 cc.

Vol. O.N. Vol. O.N. percent percent rsomerate-..-- 43. 9 97. o 49. o 97. o Extract-"---..-.-. 40. 4 109. 6 Heavy Reformate-l-Extract 35. 7 114. 4

1 Reforming Conditions:

' Weight Tempera- Hg circ.,

Hourly ture Avg., Pressure, M s.c.f./bbl.

Space F. psig. eed Velocity l2. 930 800 7, 000 il. 910 300 5, 000

Recycle, Vol. percent of fresh feed is 44 in Case I. The catalyst in both cases is SinclainBaker (platinum on ac- (solvent-'mamme giyeol with 8% by weight Hlo.)

l Demethylation (Case H, only):

Weight hourly space velocity: 1.0; Avg. Temp.: 630 F. Pla/Hydrocarbon mole ratio: 2.0; Avg. Press.: 500 p.s.i.g. Catalyst: (10% nickel molybdate on 75% Si-25% Al). Yield, Vol. percent of fresh feed: 84.5.

4 Isomerization:

weight hourly space venntyi 2;v Avg, Temp. no, reactor 760 F. :Hg/Hydrocarbon mole ratio: 3.0; Avg. Temp. 11C; reactor 725 F.

` Average reactor pressure: 5001 .-s.i.g.

Catalystf 0.6% Pd. On 87% Si-13% Al.

The intermediate vfractionation adjuncts employed to separate the eiliuent from the catalytic reformer into selected feed stocks for subsequent processing,and `the fractonator utilized to distill the raffinate from kthe sol- Vent extraction step, as well as the effluent from the catalytic demethylation, are conventional fractional distillation towers, the size of which as well as the selection 0f the operating conditions can readily be determined by those skilled in the art to which this invention pertains.

What I claim as my invention is:

l. A process for preparing high-octane number fuel from low-octane number naphtha comprising desulfurizing and vfractionating said low-octane naphtha into a higherboiling fraction and a lower boiling fraction, containing normal pentane and hexane and having an end point of about 193 F., catalytically reforming said higher Yboiling fraction in the presence of hydrogen at temperatures Vof 900-930 F. and pressures of Z50-350 lb./ sq. in., separating from the reformed product a lower boiling fraction containing benzene, toluene and xylenes admixed with parain hydrocarbons, and a higher boiling fraction, Withdrawing the higher boiling fraction from the system, solvent extracting the lower boiling reformed fraction to separate a rainate of low aromatic content and an cxtract of high aromatic content, fractionating said raiiinate into fractions, one of which contains predominantly C6 and lighter hydrocarbons, and the other of which contains C., and heavier hydrocarbons, subjecting the fraction containing C7 and heavier hydrocarbons to demethylation in the presence of hydrogen and a catalyst composed of silica-alumina cracking catalyst promoted with nickel molybdate, which is selective for the demethylation reaction at a temperature of G20-680 F. and pressure of 300-750 lb./sq. in., fractionating the demethylated product together with said ranate, combining said rst-mentioned lower boiling fraction with said raffinate fraction containing predominantly C6 and lighter hydrocarbons, separating said combined fractions into predominantly isoparaiiin and normal parain fractions and isomcrizing said normalparain fractions in the presence of a selec- 4tive noble metal isomerization catalyst at temperatures References Cited in the ile ofV this patent UNITED STATES PATENTS 2,550,531 Ciapetta Apr. 24, 1951 2,651,597 Corner et al. Sept. 8,1953 2,697,684 Hemminger et al. Dec. 21, 1954 2,769,769 Tyson Nov. 6, 1956 2,781,298 Haensel et al. Feb. 12, 1957 2,853,437 Hacnsel Spt. 23, 1958 v *Avec Le 

1. A PROCESS FOR PREPARING HIGH-OCTANE NUMBER FUEL FROM LOW-OCTANE NUMBER NAPHTHA COMPRISING DESULFURIZING AND FRACTIONATING SAID LOW-OCTANE NAPHTHA INTO A HIGHERBOILING FRACTION AND A LOWER BOILING FRACTION, CONTAINING NORMAL PENTANE AND HEXANE AND HAVING AN END POINT OF ABOUT 193* F., CATALYTICALLY REFORMING SAID HIGHER BOILING FRACTION IN THE PRESENCE OF HYDROGEN AT TEMPERATURES OF 900-930* F. AND PRESSURES OF 250-350 LB./SQ. IN., SEPARATING FROM THE REFORMED PRODUCT A LOWER BOILING FRACTION CONTAINING BENZENE, TOLUENE AND XYLENES ADMIXED WITH PARAFFIN HYDROCARBONS, AND A HIGHER BOILING FRACTION, WITH DRAWING THE HIGHER BOILING FRACTION FROM THE SYSTEM, SOLVENT EXTRACTING THE LOWER BOILING REFORMED FRACTION TO SEPARATE A RAFFINATE OF LOW AROMATIC CONTENT AND AN EXTRACT OF HIGH AROMATIC CONTENT, FRACTIONATING SAID REFFINATE INTO FRACTIONS, ONE OF WHICH CONTAINS PREDOMINANTLY C6 AND LIGHTER HYDROCARBONS, AND THE OTHER OF WHICH CONTAINS C7 AND HEAVIER HYDROCARBONS, SUBJECTING THE FRACTION CONTAINING C7 AND HEAVIER HYDROCARBONS TO DEMETHYLATION IN THE PRESENCE OF HYDROGEN AND A CATALYST COMPOSED OF SILICA-ALUMINA CRACKING CATALYST PROMOTED WITH NICKEL MILYBDATE, WHICH IS SELECTIVE FOR THE DEMETHYLATION 